Process for preparing polyolefin products

ABSTRACT

A novel liquid phase polymerization process for preparing a polyolefin product having preselected properties is disclosed. The process includes the steps of providing a liquid feedstock which contains an olefinic component and a catalyst composition consisting of a stable complex of BF 3  and a complexing agent therefor. The feedstock may comprise any one or more of a number of olefins including branched olefins such as isobutylene, C 3  to C 15  linear alpha olefins and C 4  to C 15  reactive non-alpha olefins. The feedstock and the catalyst composition are introduced into a residual reaction mixture recirculating in a loop reactor reaction zone provided in the tube side of a shell and tube heat exchanger at a recirculation rate sufficient to cause intimate intermixing of the residual reaction mixture, the added feedstock and the added catalyst composition. The heat of the polymerization reaction is removed from the recirculating intimately intermixed reaction admixture at a rate calculated to provide a substantially constant reaction temperature therein while the same is recirculating in said reaction zone. The conditions in the reactor are appropriate for causing olefinic components introduced in said feedstock to undergo polymerization to form the desired polyolefin product in the presence of the catalyst composition. A product stream containing the desired polyolefin product is withdrawn from the reaction zone. The introduction of the feedstock into the reaction zone and the withdrawal of the product stream from the reaction zone are controlled such that the residence time of the olefinic components undergoing polymerization in the reaction zone is appropriate for production of the desired polyolefin product.

CROSS REFERENCE TO RELATED APPLICATIONS

[0001] This application is a continuation of copending utilityapplication Ser. No. 09/655,084 filed on Sep. 20, 2000, which in turn isa continuation-in-part of copending utility application Ser. No.09/515,790 filed on Feb. 29, 2000. Priority from each of said copendingutility applications is claimed herein pursuant to 35 U.S.C. §120. Inaddition, priority benefits under 35 U.S.C. §119(e) are claimed in thisapplication from provisional application Serial No. 60/160,357, filed onOct. 19, 1999. The entireties of the disclosures of said priorapplications are hereby specifically incorporated herein by reference.

BACKGROUND OF THE INVENTION

[0002] 1. Field of the Invention

[0003] The present invention relates to olefin polymerization and to thepreparation of polyolefin products. In particular the present inventionrelates to the preparation of a variety of polyolefin products using aliquid phase polymerization process. In this latter regard, theinvention relates to a novel liquid phase process for the polymerizationof olefins using a modified BF₃ catalyst which is stabilized with acomplexing agent.

[0004] 2. The Prior Art Background

[0005] The polymerization of olefins using Friedel-Crafts typecatalysts, including BF₃, is a generally known procedure. The degree ofpolymerization of the products obtained varies according to which of thevarious known polymerization techniques is used. In this latter regard,it is to be understood that the molecular weight of the polymericproduct is directly related to the degree of polymerization and that thedegree of polymerization may be manipulated by manipulating processparameters so as to produce a variety of products having respectivedesired average molecular weights.

[0006] Generally speaking, due to the nature and mechanics of theolefinic polymerization process, a polyolefin product has a singledouble bond remaining in each molecule at the end of the polymerizationprocess. The position of this remaining double bond is often animportant feature of the product. For example, polyisobutylene (PIB)molecules wherein the remaining double bond is in a terminal(vinylidene) position are known to be more reactive than PIB moleculeswherein the remaining double bond is internal, that is, not in aterminal position. A PIB product wherein at least 50% of the doublebonds are in a terminal position may often be referred to as highvinylidene or highly reactive PIB. The extent to which a polyolefinproduct has terminal double bonds may also be manipulated bymanipulation of process parameters.

[0007] Current processes for olefin oligomerization often employBF₃/co-catalyst systems wherein the BF₃ is complexed with a co-catalyst.This is done for a variety of reasons that are well known to thoseskilled in the olefin polymerization field. For example, and as isexplained in U.S. Pat. No. 5,408,018, a complexed BF₃ catalyst may beuseful for manipulating and attempting to balance the molecular weight,vinylidene content and polydispersity of PIB. The co-catalyst often ispropanol or a higher alcohol and such co-catalyst systems are usedirrespective of whether the desired product is a poly alpha olefin or apoly internal olefin. However, the use of alcohols having beta hydrogenatoms in such co-catalyst complexes is troublesome because, over time,the BF₃ tends to attack the beta hydrogen atoms. This leads todecomposition of the alcohol whereby the catalyst is renderedineffective. Thus, the co-catalyst complex is unstable and often has avery short shelf life.

[0008] To address this problem, many current processes employ aprocedure whereby the co-catalyst complex is prepared in-situ by mixingthe alcohol and gaseous BF₃ immediately prior to introduction of theco-catalyst complex into a reactor. In addition, it is not unusual inthe conduct of processes employing such co-catalyst systems to use anexcess of alcohol and to sparge gaseous BF₃ into the reaction mass atseveral downstream points to rebuild catalyst activity. Such methodologyimplies a three-phase reaction and the necessity of using a stirred tankreactor to provide means of dispersing gaseous BF₃ into the reactionmass. These processes use either batch reactors or a set of continuouslystirred tank reactors in series to provide both gas handling capabilityand to satisfy the necessity for a plug flow reactor configuration.

[0009] It is also known that alpha olefins, particularly PIB, may bemanufactured in at least two different grades—regular and highvinylidene. Conventionally, these two product grades have been made bydifferent processes, but both often and commonly use a diluted feedstockin which the isobutylene concentration may range from 40-60% by weight.More recently it has been noted that at least the high vinylidene PIBmay be produced using a concentrated feedstock having an isobutylenecontent of 90% by weight or more. Non-reactive hydrocarbons, such asisobutane, n-butane and/or other lower alkanes commonly present inpetroleum fractions, may also be included in the feedstock as diluents.The feedstock often may also contain small quantities of otherunsaturated hydrocarbons such as 1-butene and 2-butene.

[0010] Regular grade PIB may range in molecular weight from 500 to1,000,000 or higher, and is generally prepared in a batch process at lowtemperature, sometimes as low as −50 to −70° C. AlCl₃, RAlCl₂ or R₂AlClare used as catalysts. The catalyst is not totally removed from thefinal PIB product. Molecular weight may be controlled by temperaturesince the molecular weight of the product varies inversely withtemperature. That is to say, higher temperatures give lower molecularweights. Reaction times are often in the order of hours. The desiredpolymeric product has a single double bond per molecule, and the doublebonds are mostly internal. Generally speaking, at least about 90% of thedouble bonds are internal and less than 10% of the double bonds are in aterminal position. Even though the formation of terminal double bonds isbelieved to be kinetically favored, the long reaction times and the factthat the catalyst is not totally removed, both favor the rearrangementof the molecule so that the more thermodynamically favored internaldouble bond isomers are formed. Regular PIB may be used as a viscositymodifier, particularly in lube oils, as a thickener, and as a tackifierfor plastic films and adhesives. PIB can also be functionalized toproduce intermediates for the manufacture of detergents and dispersantsfor fuels and lube oils.

[0011] High vinylidene PIB, a relatively new product in the marketplace,is characterized by a large percentage of terminal double bonds,typically greater than 70% and preferentially greater than 80%. Thisprovides a more reactive product, compared to regular PIB, and hencethis product is also referred to as highly reactive PIB. The termshighly reactive (HR-PIB) and high vinylidene (HV-PIB) are synonymous.The basic processes for producing HV-PIB all include a reactor system,employing BF₃ and/or modified BF₃ catalysts, such that the reaction timecan be closely controlled and the catalyst can be immediatelyneutralized once the desired product has been formed. Since formation ofthe terminal double bond is kinetically favored, short reactions timesfavor high vinylidene levels. The reaction is quenched, usually with anaqueous base solution, such as, for example, NH₄OH, before significantisomerization to internal double bonds can take place. Molecular weightsare relatively low. HV-PIB having an average molecular weight of about950-1050 is the most common product. Conversions, based on isobutylene,are kept at 75-85%, since attempting to drive the reaction to higherconversions reduces the vinylidene content through isomerization. PriorU.S. Pat. No. 4,152,499 dated May 1, 1979, U.S. Pat. No. 4,605,808 datedAug. 12, 1986, U.S. Pat. No. 5,068,490 dated Nov. 26, 1991, U.S. Pat.No. 5,191,044 dated Mar. 2, 1993, U.S. Pat. No. 5,286,823 dated Jun. 22,1992, U.S. Pat. No. 5,408,018 dated Apr. 18, 1995 and U.S. Pat. No.5,962,604 dated Oct. 5, 1999 are all directed to related subject matter.

[0012] U.S. Pat. No. 4,152,499 describes a process for the preparationof PIBs from isobutylene under a blanket of gaseous BF₃ acting as apolymerization catalyst. The process results in the production of a PIBwherein 60 to 90% of the double bonds are in a terminal (vinylidene)position.

[0013] U.S. Pat. No. 4,605,808 discloses a process for preparing PIBwherein a catalyst consisting of a complex of BF₃ and an alcohol isemployed. It is suggested that the use of such a catalyst complexenables more effective control of the reaction parameters. Reactioncontact times of at least 8 minutes are required to obtain a PIB productwherein at least about 70% of the double bonds are in a terminalposition.

[0014] U.S. Pat. No. 5,191,044 discloses a PIB production processrequiring careful pretreatment of a BF₃/alcohol complex to insure thatall free BF₃ is absent from the reactor. The complex must contain asurplus of the alcohol complexing agent in order to obtain a productwherein at least about 70% of the double bonds are in a terminalposition. The only reaction time exemplified is 10 minutes, and thereaction is carried out at temperatures below 0° C.

[0015] In addition to close control of reaction time, the key toobtaining high vinylidene levels seems to be control of catalystreactivity. This has been done in the past by complexing BF₃ withvarious oxygenates including sec-butanol and MTBE. One theory is thatthese complexes are actually less reactive than BF₃ itself,disproportionately slowing the isomerization reaction and thus allowingfor greater differentiation between the vinylidene forming reaction(polymerization) and the isomerization reaction rates. Mechanisms havealso been proposed that suggest the BF₃ complexes are non-protonated andthus are not capable of isomerizing the terminal double bond. Thisfurther suggests that water (which can preferentially protonate BF₃)must generally be excluded from these reaction systems. In fact, priorpublications describing preparation of PIB using BF₃ complexes teach lowwater feed (less than 20 ppm) is critical to formation of the highvinylidene product.

[0016] HV-PIB is increasingly replacing regular grade PIB for themanufacture of intermediates, not only because of higher reactivity, butalso because of developing requirements for “chloride free” materials inthe final product applications. Important PIB derivatives are PIBamines, PIB alkylates and PIB maleic anhydride adducts.

[0017] PIB amines can be produced using a variety of proceduresinvolving different PIB intermediates which provide a reactive site forsubsequent amination. These intermediates may include, for example,epoxides, halides, maleic anhydride adducts, and carbonyl derivatives.

[0018] Reference to HV-PIB as “highly reactive” is relative to regulargrade PIB. HV-PIB is still not, in absolute terms, highly reactivetoward formation of some of these intermediates. Other classes ofcompounds, polyethers for example, can be much more reactive in theformation of amines and amine intermediates. Amines derived frompolyethers are known as polyether amines (PEAs) and are competitiveproducts to PIB amines.

[0019] The use of HV-PIB as an alklylating agent for phenolic compounds,is triggered by the higher reactivity and higher yields achievable withHV-PIB. These very long chain alkyl phenols are good hydrophobes forsurfactants and similar products.

[0020] The largest volume PIB derivatives are the PIB-maleic anhydridereaction products. HV-PIB is reacted with maleic anhydride through thedouble bond giving a product with anhydride functionality. Thisfunctionality provides reactivity for the formation of amides and othercarboxylate derivatives. These products are the basis for most of thelube oil detergents and dispersants manufactured today. As mentionedabove, PIB-maleic anhydride products can also be used as intermediatesin the manufacture of PIB amine fuel additives.

[0021] Other polyolefins which are commercially useful for a variety ofpurposes include conventional PIB wherein the vinylidene content is lessthan 50%, low molecular weight (<350 and perhaps even <250) oligomers ofbranched monomers such as isobutylene, oligomers and higher molecularweight polymers of linear C₃-C₁₅ alpha olefins, and oligomers and highermolecular weight polymers of linear C₄-C₁₅ non-alpha (internal doublebond) olefins. Although these materials are all well known to thoseskilled in the olefin polymerization field, there is always a need fornew developments which improve process efficiency and/or productqualities and reduce operating costs and/or capital expenditures.

SUMMARY OF THE INVENTION

[0022] The present invention provides a novel process for the efficientand economical production of polyolefin products. Generally speaking,the invention provides a liquid phase polymerization process forpreparing a polyolefin product having preselected properties. Inaccordance with the principles and concepts of the invention, theprocess includes the steps of providing a liquid feedstock comprising atleast one olefinic component and a catalyst composition comprising astable complex of BF₃ and a complexing agent therefor. The feedstock andthe catalyst composition are introduced into a residual reaction mixturein a loop reactor reaction zone where the residual reaction mixture isrecirculated at a recirculation rate sufficient to cause intimateintermixing of the residual reaction mixture, the added feedstock andthe added catalyst composition to thereby present a recirculating,intimately intermixed reaction admixture in said reaction zone. Therecirculating intimately intermixed reaction admixture is maintained inits intimately intermixed condition while the heat of reaction isremoved therefrom at a rate calculated to provide a substantiallyconstant reaction temperature in the reaction admixture while the sameis recirculating in said reaction zone. The constant reactiontemperature is at a level appropriate for causing olefinic componentsintroduced in said feedstock to undergo polymerization to form thedesired polyolefin product in the presence of the catalyst composition.A product stream comprising the desired polyolefin product is withdrawnfrom the reaction zone. In accordance with the invention, theintroduction of the feedstock into the reaction zone and the withdrawalof the product stream from the reaction zone are controlled such thatthe residence time of the olefinic components undergoing polymerizationin the reaction zone is appropriate for production of the desiredpolyolefin product.

[0023] In accordance with one preferred form of the invention, thereaction zone may comprises the tube side of a shell-and-tube heatexchanger. The heat of the exothermic olefin polymerization reaction maybe removed simultaneously with its generation by circulation of acoolant in the shell side of the exchanger. Preferably, the residencetime of the olefinic components undergoing polymerization may be nogreater than about 3 minutes. Even more preferably, such residence timemay be no greater than about 2 minutes. More preferably still, suchresidence time may be no greater than about 1 minute. Ideally, theresidence time may be less than 1 minute.

[0024] In accordance with another preferred form of the invention, thecomplexing agent should preferably be such the a stable catalyst complexis formed with BF₃. This is particularly advantageous at the relativelyhigh reaction temperatures needed for oligomerization processes. In thisregard, the complexing agent may advantageously comprise an alcohol,preferably a primary alcohol, and even more preferably a C₁-C₈ primaryalcohol. In a highly preferred form of the invention, the alcohol shouldhave no hydrogen atom on a β carbon. In this highly preferred form ofthe invention, the alcohol may be, for example, methanol or neopentanol.

[0025] In accordance with yet another preferred form of the invention,the complexing agent may comprise a glycol, preferably glycol whereineach hydroxyl group of the glycol is in a primary position, and evenmore preferably a C₁-C₈ glycol wherein each hydroxyl group of the glycolis in a primary position. In this highly preferred form of theinvention, the glycol may be, for example, ethylene glycol.

[0026] In conformity with the concepts and principles of another aspectof the invention, the molar ratio of BF₃ to complexing agent in catalystcomplex may range from approximately 0.5:1 to approximately 5:1.Preferably the molar ratio of BF₃ to complexing agent in said complexmay range from approximately 0.5:1 to approximately 2:1. Even morepreferably, the molar ratio of BF₃ to complexing agent in the complexmay range from approximately 0.5:1 to approximately 1:1. Ideally, themolar ratio of BF₃ to complexing agent in complex may be approximately1:1. Alternatively, the molar ratio of BF₃ to complexing agent in saidcomplex may be approximately 0.75:1.

[0027] According to another aspect of the invention, the process maydesirably be conducted such that from about 0.1 to about 10 millimolesof BF₃ are introduced into the reaction admixture with said catalystcomposition for each mole of olefinic component introduced into saidadmixture in said feedstock. Preferably, from about 0.5 to about 2millimoles of BF₃ may be introduced into the reaction admixture with thecatalyst composition for each mole of olefinic component introduced intothe admixture in said feedstock.

[0028] Another important preferred feature of the invention involves thecontinuous recirculation of the reaction admixture at a first volumetricflow rate, and the continuous introduction of the feedstock and thecatalyst composition at a combined second volumetric flow rate.Desirably the ratio of the first volumetric flow rate to the secondvolumetric flow rate may range from about 20:1 to about 50:1. Preferablythe ratio of the first volumetric flow rate to the second volumetricflow rate may range from about 25:1 to about 40:1. Ideally the ratio ofthe first volumetric flow rate to the second volumetric flow rate mayrange from about 28:1 to about 35:1. With regard to this latter aspectof the invention, the ratio of the first volumetric flow rate to thesecond volumetric flow rate may be such that the concentrations ofingredients in the reaction admixture remain essentially constant andsuch that essentially isothermal conditions are established andmaintained in said reaction admixture.

[0029] In conformity with the principles and concepts of the invention,the feedstock and the catalyst composition may be premixed andintroduced into the reaction zone together as a single stream at saidsecond volumetric flow rate. Alternatively, the feedstock and thecatalyst composition may be introduced into the reaction zone separatelyas two streams, the flow rates of which together add up to said secondvolumetric flow rate.

[0030] In further conformity with the principles and concepts of theinvention, the reactor configuration, the properties of the reactionmixture, and the first volumetric flow rate may preferably be such thatturbulent flow is maintained in said reaction zone. In this regard, inan ideal form of the invention, a Reynolds number of at least about 2000is maintained in said reaction zone. In still further conformity withthe principles and concepts of the invention, the reactor may take theform of the tube side of a shell-and-tube heat exchanger. In thisregard, in an ideal form of the invention, a U of at least about 50Btu/min ft² ° F. is maintained in reaction zone.

[0031] Preferably, in accordance with the invention, the feed stock maycomprise at least about 30% by weight of said olefinic component.Additionally, the feed stock may include non-reactive hydrocarbondiluents. In this latter regard, the feed stock may comprise at leastabout 30% by weight of said olefinic component with the remainder beingnon-reactive hydrocarbon diluents.

[0032] The polymerization process of the invention may be a cationicprocess. Alternatively the polymerization process of the invention maybe a covalent process. An important feature of the invention is that thepolyolefin product of the process of the invention may have a molecularweight of at least about 350 but no more than about 5000. Alternatively,the polyolefin product of the process of the invention may have amolecular weight no greater than about 350 and perhaps no greater thanabout 250.

[0033] In accordance with an important aspect of the invention, theolefinic component which is subjected to polymerization may compriseisobutylene and the polyolefin product may comprise PIB. In furtheraccordance with this aspect of the invention, the PIB may have avinylidene content of at least about 50%. Alternatively, the PIB mayhave a vinylidene content no greater than about 50%.

[0034] In accordance with yet another important aspect of the invention,the olefinic component may be a branched compound and the product maycomprise a two, three or four member oligomer. The olefinic componentused in the process of the invention may comprise isobutylene and thepolyolefin product may comprise a C₁₂, C₁₆, C₂₀, or C₂₄ PIB oligomer.Alternatively, the olefinic component may comprise either a C₃ to C₁₅linear alpha olefin or a C₄ to C₁₅ reactive non-alpha olefin such as2-butene.

[0035] The present invention further provides a novel process for theefficient and economical production of HV-PIB. Generally speaking, theinvention provides a HV-PIB production process wherein thepolymerization reaction takes place at higher temperatures and at lowerreaction times than were thought possible in the past. In particular,the present invention provides a liquid phase polymerization process forpreparing low molecular weight, highly reactive polyisobutylene.Generally speaking, the process may involve cationic polymerization.However, under some conditions the polymerization reaction may becovalent. Particularly the latter may be true when ether is used as acomplexing agent. In accordance with this embodiment of the invention,the process includes the provision of a feedstock comprising isobutyleneand a catalyst composition comprising a complex of BF₃ and a complexingagent. The feedstock and the catalyst composition are introduced eitherseparately or as a single mixed stream into a residual reaction mixturein a reaction zone. The residual reaction mixture, the feedstock and thecatalyst composition are then intimately intermixed to present anintimately intermixed reaction admixture in said reaction zone. Thereaction admixture is maintained in its intimately intermixed conditionand kept at a temperature of at least about 0° C. while the same is insaid reaction zone, whereby the isobutylene in the reaction admixture iscaused to undergo polymerization to form a polyisobutylene product. Aproduct stream comprising a low molecular weight, highly reactivepolyisobutylene is then withdrawn from the reaction zone. Theintroduction of the feedstock into said reaction zone and the withdrawalof the product stream from the reaction zone are controlled such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is no greater than about 4 minutes. In accordance with theinvention, it is possible to conduct the reaction so that the residencetime is no greater than about 3 minutes, no greater than about 2minutes, no greater than about 1 minute, and ideally, even less than 1minute.

[0036] In accordance with the concepts and principles of the invention,the process may be conducted in a manner such that the polyisobutylenethus produced has a molecular weight in the range of from about 250 toabout 5000, in the range of from about 600 to about 4000, in the rangeof from about 700 to about 3000, in the range of from about 800 to about2000, and ideally in the range of from about 950 to about 1050. Inaccordance with the invention, it is possible to so control the processthat a particular molecular weight, such as for example, a molecularweight of about 1000, may be achieved.

[0037] A major purpose of the invention is to provide a process whichmay be controlled sufficiently to insure the production of apolyisobutylene product having a vinylidene content of at least about70%. More preferably the PIB product may have a vinylidene content of atleast about 80%. Vinylidene contents of at least about 90% may also beachieved through the use of the invention.

[0038] The complexing agent used to complex with the BF₃ catalyst maydesirably be an alcohol, and preferably may be a primary alcohol. Morepreferably the complexing agent may comprise a C₁-C₈ primary alcohol andideally may be methanol.

[0039] To achieve the desired results of the invention, the molar ratioof BF₃ to complexing agent in the complex may range from approximately0.5:1 to approximately 5:1. Preferably the molar ratio of BF₃ tocomplexing agent in the complex may range from approximately 0.5:1 toapproximately 2:1. Even more preferably the molar ratio of BF₃ tocomplexing agent in the complex may range from approximately 0.5:1 toapproximately 1:1, and ideally, the molar ratio of BF₃ to complexingagent in the complex may be approximately 1:1.

[0040] According to the principles and concepts of the invention, it ispreferred that from about 0.1 to about 10 millimoles of BF₃ may beintroduced into the reaction admixture with the catalyst composition foreach mole of isobutylene introduced into the admixture in the feedstock.Even more preferably, from about 0.5 to about 2 millimoles of BF₃ may beintroduced into the reaction admixture with said catalyst compositionfor each mole of isobutylene introduced into the admixture in thefeedstock.

[0041] The invention provides a process whereby the polydispersity ofsaid polyisobutylene may be no more than about 2.0, and desirably may beno more than about 1.65. Ideally, the polydispersity may be in the rangeof from about 1.3 to about 1.5.

[0042] In accordance with one preferred aspect of the invention, thereaction zone may comprise a loop reactor wherein the reaction admixtureis continuously recirculated at a first volumetric flow rate, and saidfeedstock and said catalyst composition are continuously introduced at acombined second volumetric flow rate. The ratio of said first volumetricflow rate to said second volumetric flow rate may desirably range fromabout 20:1 to about 50:1, may preferably range from about 25:1 to about40:1 and ideally may range from about 28:1 to about 35:1. In order toachieve the benefits of the invention, the ratio of said firstvolumetric flow rate to said second volumetric flow rate may preferablybe such that the concentrations of ingredients in the reaction admixtureremain essentially constant and/or such that essentially isothermalconditions are established and maintained in said reaction admixture.

[0043] The feedstock and the catalyst composition may be premixed andintroduced into the reaction zone together as a single stream at saidsecond volumetric flow rate. Alternatively, the feedstock and thecatalyst composition may be introduced into the reaction zone separatelyas two respective streams, the flow rates of which together add up tosaid second volumetric flow rate.

[0044] To achieve the desired results of the invention, the reactorconfiguration, the properties of the reaction mixture, and the firstvolumetric flow rate may be such that turbulent flow is maintained insaid reaction zone. In particular, the system may be such that aReynolds number of at least about 2000 is achieved and maintained insaid reaction zone. The system may also be such that a heat transfercoefficient (U) of at least about 50 Btu/min ft² ° F. is achieved andmaintained in said reaction zone. To this end, the reactor maypreferably be the tube side of a shell-and-tube heat exchanger.

[0045] In further accordance with the concepts and principles of theinvention, the feed stock may generally comprise at least about 30% byweight of isobutylene, with the remainder being non-reactive hydrocarbondiluents.

[0046] In a more specific sense, the invention may provide a liquidphase polymerization process for preparing polyisobutylene having anaverage molecular weight in the range of from about 500 to about 5000and a vinylidene content of at least 70%. The process may compriseproviding both a feedstock comprising isobutylene and a separatecatalyst composition made up of a complex of BF₃ and a C₁ to C₈ primaryalcohol. The molar ratio of BF₃ to alcohol in said complex may desirablybe in the range of from about 0.5:1 to about 2:1. The feedstock and thecatalyst composition may be introduced separately or together as asingle stream into a residual reaction mixture in a reaction zone, andthe residual reaction mixture, the feedstock and the catalystcomposition may be intimately intermixed to present an intimatelyintermixed reaction admixture in said reaction zone. The introduction ofthe catalyst complex into the reaction admixture may preferably becontrolled so that about 0.1 to about 10 millimoles of BF₃ areintroduced for each mole of isobutylene introduced with the feedstock.The intimately intermixed condition of the reaction admixture shouldpreferably be maintained and the temperature thereof kept at about 0° C.or above while the admixture is in the reaction zone, whereby theisobutylene in the admixture undergoes polymerization to form saidpolyisobutylene. Thereafter, a product stream comprising thepolyisobutylene product may be withdrawn from the reaction zone. Theintroduction of said feedstock into the reaction zone and the withdrawalof the product stream from the reaction zone may preferably be such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is no greater than about 4 minutes.

[0047] Even more desirably, the invention may provide a liquid phasepolymerization process for preparing polyisobutylene having an averagemolecular weight in the range of from about 950 to about 1050, apolydispersity within the range of from about 1.3 to about 1.5, and avinylidene content of at least about 80%. In accordance with thispreferred aspect of the invention, the process comprises providing botha feedstock made up of at least about 40% by weight isobutylene and aseparate catalyst composition made up of a complex of BF₃ and methanol,wherein the molar ratio of BF₃ to methanol in the complex ranges fromabout 0.5:1 to about 1:1. The feedstock and the catalyst composition areintroduced either separately or together into a residual reactionmixture in a reaction zone. The residual reaction mixture, the feedstockand the catalyst composition are intimately intermixed by turbulent flowwithin said reaction zone, whereby an intimately intermixed reactionadmixture is present in the reaction zone. Preferably, the catalystcomplex is introduced into the reaction admixture at a rate such thatabout 0.5 to about 2 millimoles of BF₃ are introduced for each mole ofisobutylene introduced in the feedstock. The intimately intermixedcondition of the reaction admixture is maintained and the temperaturethereof is kept at about 0° C. or more while the same is in saidreaction zone, whereby the isobutylene therein is caused to undergopolymerization to form said polyisobutylene. A product stream comprisingsaid polyisobutylene is withdrawn from said reaction zone. In accordancewith the invention, the introduction of feedstock into the reaction zoneand the withdrawal of product stream therefrom are controlled such thatthe residence time of the isobutylene undergoing polymerization in thereaction zone is within the range of from about 45 to about 90 seconds.

BRIEF DESCRIPTION OF THE DRAWING

[0048]FIG. 1 is a schematic illustration of a reactor in the form of amulti-pass shell and tube heat exchanger which is useful for carryingout the improved process of the invention; and

[0049]FIG. 2 is a schematic illustration of an alternative reactor inthe form of a single pass shell and tube exchanger which is also usefulfor carrying out the improved process of the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0050] In accordance with one very important embodiment of the presentinvention, an improved liquid phase process is provided for theefficient and economical production of PIB. In accordance with thisembodiment of the invention, an isobutylene containing feedstock streamis contacted in a reaction zone with a catalyst which facilitates thepolymerization reaction. Appropriate reaction conditions are provided inthe reaction zone. After an appropriate residence time, a PIB containingproduct stream is withdrawn from the reaction zone. As mentioned above,many techniques for conducting the reaction are known; however, from acommercial viewpoint it is always desirable to improve the efficiencyand economics of the process. With the foregoing in mind, the presentinvention provides an improved PIB producing process which may be easilycontrolled and manipulated to efficiently and economically provide arelatively low molecular weight, highly reactive PIB product.

[0051] The improved process of the present invention features the use ofa BF₃ catalyst which desirably may be complexed with a complexing agentwhich appropriately alters the performance of the catalyst. Many otherpotentially useful catalysts are known to those of ordinary skill in therelated art field. In particular, many useful catalysts are described inthe prior patents referenced above. The complexing agent for thecatalyst, and in particular for the BF₃ catalyst, may be any compoundcontaining a lone pair of electrons, such as, for example, an alcohol,an ester or an amine. For purposes of the present invention, however,the complexing agent preferably may be an alcohol, desirably a primaryalcohol, preferably a C₁-C₈ primary alcohol and ideally methanol.

[0052] The molar ratio of BF₃ to complexing agent in the catalystcomposition may generally be within the range of from approximately0.5:1 to approximately 5:1, desirably within the range of fromapproximately 0.5:1 to approximately 2:1, and preferably within therange of from approximately 0.5:1 to approximately 1:1. Ideally, thecatalyst composition may simply be a 1:1 complex of BF₃ and methanol. Insome preferred embodiments of the invention, the molar ratio of BF₃ tocomplexing agent in said complex may be approximately 0.75:1.

[0053] The temperature in the reaction zone may generally and preferablybe greater than 0° C., the reactor residence time may generally andpreferably be less than 4 minutes and the desired vinylidene (terminalunsaturation) content in the PIB product may preferably and generally begreater than about 70%. With these parameters, it is possible to operatethe process so as to achieve efficiencies and economies not previouslythought to be available. In accordance with the present invention, thecatalyst concentration and the BF₃/complexing ratio may be manipulatedas required to achieve the desired 70% vinylidene content with areaction temperature greater than 0° C. and a reactor residence time ofless than 4 minutes. Generally speaking, for PIB production the amountof the BF₃ catalyst introduced into the reaction zone should be withinthe range of from about 0.1 to about 10 millimoles for each mole ofisobutylene introduced into the reaction zone. Preferably, the BF₃catalyst may be introduced at a rate of about 0.5 to about 2 millimolesper mole of isobutylene introduced in the feedstock.

[0054] The process itself includes steps resulting in the intimatemixing of the isobutylene containing reactant stream and the catalystcomplex and/or removal of heat during the reaction. The intimate mixingmay desirably be accomplished by turbulent flow. Turbulent flow alsoenhances heat removal. These conditions separately or together permitthe higher operating temperatures (e.g. >0° C.) and the shorter reactorresidence times (e.g. <4 minutes) provided by the invention. Theseimportant parameters may be achieved by causing the catalyzed reactionto take place in the tubes of a shell-and-tube heat exchanger at a flowrate which results in turbulent flow.

[0055] Many potentially valuable reactors are well known to theroutineers in the art to which the invention pertains. However, forpurposes of one preferred embodiment of the invention, the reactor maybe a four-pass shell-and-tube heat exchanger as shown in FIG. 1 where itis identified by the numeral 10. The reactor may, for example, have 80⅜-inch tubes with a wall thickness of 0.022 inch, each thereby providingan internal tube diameter of 0.331 inch. The reactor may be three feetlong and may have internal baffling and partitions to provide 4 passeswith 20 tubes per pass. Such construction is well known in the heatexchanger and reactor arts and no further explanation is believednecessary.

[0056] In operation, in accordance with the preferred procedure forproducing highly reactive PIB, the isobutylene containing feedstockenters the reactor system through pipe 15 which is preferably locatedadjacent the bottom head 11 of reactor 10. Pipe 15 directs the feedstock into the suction line 20 of a recirculation pump 25. The catalystcomplex may be injected into the reactor circulation system through pipe30 located adjacent bottom head 11 of reactor 10. It should be notedhere, that in accordance with the principles and concepts of theinvention, the catalyst complex could just as well be injectedseparately into the reactor, in which case a separate catalyst pumpmight be required.

[0057] A catalyst modifier may be added to the feedstock via pipe 16before the feedstock enters the reactor system. The purpose of themodifier is to assist in controlling the vinylidene content of the PIBproduct. The catalyst modifier may be any compound containing a lonepair of electrons such as an alcohol, an ester or an amine. However, itis pointed out in this regard that if the amount of modifier is toogreat, the same may actually kill the catalyst. The feedstock containingthe modifier enters the reactor system at the suction line 20 of thecirculation pump 25. The catalyst complex composition enters the reactorsystem via line 30 at a location downstream from pump 25 and adjacentthe first pass as shown in FIG. 1. The catalyst complex is preferably amethanol/BF₃ complex with a 1:1 molar ratio of BF₃ to methanol. Theamount of modifier added via line 16 may vary from 0 to about 1 mole foreach mole of BF₂ added as a complex via line 30.

[0058] Circulation pump 25 pushes the reaction mixture through line 35,control valve 40 and line 45 into the bottom head 11 of the reactor 10.A flow meter 46 may be positioned in line 45 as shown. The reactionmixture travels upwardly through pass 50, downwardly through pass 51,upwardly through pass 52 and downwardly through pass 53. As explainedpreviously, each pass 50, 51, 52 and 53 may preferably include 20separate tubes. For clarity, only a respective single tube isillustrated schematically in each pass in FIG. 1. These tubes areidentified by the reference numerals 50 a, 51 a, 52 a and 53 a. However,as discussed above, each pass may preferably consist of a plurality, forexample, 20 of these individual tubes.

[0059] It is to be noted here, that the reaction mixture shouldpreferably be circulated through the tubes 50 a, 51 a, 52 a, 53 a of thereactor at a flow rate sufficient to obtain turbulent flow, whereby toachieve intimate intermixing between the catalyst complex and thereactants and a heat transfer coefficient appropriate to provide propercooling. In this regard, the flow rate, the reaction mixture properties,the reaction conditions and the reactor configuration should beappropriate to produce a Reynolds number (Re) in the range of from about2000 to about 3000 and a heat transfer coefficient (U) in the range offrom about 50 to about 150 Btu/min ft² ° F. in the tubes of the reactor.Such parameters may generally be obtained when the linear flow rate of atypical reaction mixture through a tube having an internal diameter of0.331 inch is within the range of from about 6 to 9 feet per second.

[0060] The circulating reaction mixture leaves reactor 10 via suctionline 20. The recirculation rate in the system is preferably sufficientlyhigh so that the reactor, in essence, is a Continuous Stirred TankReactor (CSTR). In this same connection, and generally speaking, therecirculation rate of the reaction mixture should preferably be suchthat essentially steady state conditions are maintained in the reactor.It is pointed out in this latter regard that regardless of the systembeing in an unsteady or steady state, the design equations may bereduced to those of a CSTR when the recycle rate is sufficiently high.The reactor may also be of the type which is sometimes referred to as aloop reactor. With this system, which is only a preferred system sincethere are many other arrangements which would be apparent to those ofordinary skill in the art, the flow rate of the reactant mixture in thereactor may be adjusted and optimized independently of feed stockintroduction and product removal rates so as to achieve thoroughintermixing of the catalyst complex and the reactants and appropriatetemperature control.

[0061] A product exit line 55 may preferably be provided in top head 12at a point which is approximately adjacent the transition zone betweenthe third and fourth passes. Such positioning may be desirable to avoidany potential for loss of unreacted isobutylene. Moreover, thepositioning of the exit line 55 should be appropriate to facilitatebleeding of gas from the reactor during startup. A coolant may desirablybe circulated on the shell side of the reactor at a rate to remove heatof reaction and maintain the desired temperature in the reactor.

[0062] The product exiting the system via line 55 should be quenchedimmediately with a material capable of killing the catalyst, such as,for example, ammonium hydroxide. Thus, any potential rearrangement ofthe polymer molecule which would shift the double bond away from theterminal position is minimized. The high vinylidene isobutylene productmay then be directed to a work up system (not shown) where catalystsalts may be removed and the isobutylene product separated fromunreacted isobutylene and other undesirable contaminants such asdiluents, etc. These latter materials may then be recycled or divertedfor other uses employing known methodology.

[0063] With the described recirculation system, the rate of feedstockintroduction into the reaction mixture and the rate of product removalare each independent of the circulation rate. As will be appreciated bythose of ordinary skill in the art, the number of passes through thereactor and the size and configuration of the latter are simply mattersof choice. The feedstock and product withdrawal flow rates maypreferably be chosen such that the residence time of the reactionmixture within the reactor is 4 minutes or less, desirably 3 minutes orless, preferably 2 minutes or less, even more preferably 1 minute orless, and ideally less than 1 minute. From a commercial operatingviewpoint, the flow rate should be such that the residence time of thereaction mixture in the reactor is within the range of from about 45 toabout 90 seconds. In connection with the foregoing, the residence timeis defined as the total reactor system volume divided by the volumetricflow rate.

[0064] The recirculation flow rate, that is the flow rate of thereaction mixture in the system induced by the recirculation pump 25, iscontrolled, as described above, to achieve appropriate turbulence and/orheat transfer characteristics. This recirculation flow rate is often afunction of the system itself and other desired process conditions. Forthe system described above, the ratio of the recirculation flow rate tothe incoming feedstock flow rate (recycle ratio) should generally bemaintained in the range of from about 20:1 to about 50:1, desirably inthe range of from about 25:1 to about 40:1, and ideally in the range offrom about 28:1 to about 35:1. In particular, in addition to causingturbulence and providing an appropriate heat transfer coefficient, therecirculation flow rate of the reaction mixture should be sufficient tokeep the concentrations of the ingredients therein essentially constantand/or to minimize temperature gradients within the circulating reactionmixture whereby essentially isothermal conditions are established andmaintained in the reactor.

[0065] As mentioned above, the recycle ratios generally may be in therange of from about 20:1 to about 50:1 when the desired product ishighly reactive PIB. Higher recycle ratios increase the degree of mixingand the reactor approaches isothermal operation leading to narrowerpolymer distributions. Lower recycle ratios decrease the amount ofmixing in the reactor, and as a result, there is a greater discrepancyin the temperature profiles. As the recycle ratio approaches zero, thedesign equations for the reactor reduce to those for a plug flow reactormodel. On the other hand, as the recycle ratio approaches infinity, themodeling equations reduce to those for a CSTR. When CSTR conditions areachieved, both temperature and composition remain constant and thecomposition of the product stream leaving the reactor is identical tothe composition of the reaction mixture recirculating in the reactor.

[0066] Needless to say, after steady state or near steady stateoperation has been established in the reactor, as the feedstock entersthe system, an equal volume of product is pushed out of the reactorloop. Under CSTR conditions, the point at which the product stream iswithdrawn is independent of reactor geometry. However, the top of thethird pass was chosen for this particular embodiment of the invention soany air or non-condensable species in the reactor at start-up mayconveniently be purged. Also, it is preferred that the withdrawal pointbe as far as possible from the point where fresh feedstock is introducedinto the system just to make sure that conditions within the reactorhave achieved steady state operation and are therefore as stable aspossible.

[0067] When highly reactive PIB is the desired product, the feedstockentering the system through line 15 may be any isobutylene containingstream such as, but not limited to, isobutylene concentrate, dehydroeffluent, or a typical raff-1 stream. These materials are describedrespectively below in Tables 1, 2 and 3. TABLE 1 Isobutylene ConcentrateIngredient Weight % C₃ component 0.00 I-butane 6.41 n-butane 1.681-butene 1.30 I-butene 89.19 trans-2-butene 0.83 cis-2-butene 0.381,3-butadiene 0.21

[0068] TABLE 2 Dehydro Effluent Ingredient Weight % C₃ components 0.38I-butane 43.07 n-butane 1.29 1-butene 0.81 I-butene 52.58 trans-2-butene0.98 cis-2-butene 0.69 1,3-butadiene 0.20

[0069] TABLE 3 Raff-1 Ingredient Weight % C₃ components 0.57 I-butane4.42 n-butane 16.15 1-butene 37.22 I-butene 30.01 trans-2-butene 8.38cis-2-butene 2.27 1,3-butadiene 0.37 MTBE 0.61

[0070] For commercial and process economies, the isobutylene content ofthe feedstock generally should be at least about 30 weight %, with theremainder comprising one or more non-reactive hydrocarbon, preferablyalkane, diluents.

[0071] The desired product is a relatively low molecular weight, highlyreactive polyisobutylene. Thus, the polyisobutylene leaving the reactorby way of line 55 should have an average molecular weight which is lessthan about 10,000. Generally speaking, the produced isobutylene shouldhave an average molecular weight within the range of from about 500 toabout 5000, desirably from about 600 to about 4000, preferably fromabout 700 to about 3000, even more preferably from about 800 to about2000, and ideally from about 950 to about 1050. By carefully controllingthe various parameters of the process, it might even be possible toproduce a product wherein the average molecular weight is consistent atsome desired number, for example, 1000.

[0072] The polydispersity of the PIB may also be important. The termpolydispersity refers to the molecular weight distribution in a givenpolymer product and generally is defined as the ratio of the molecularweight of the highest molecular weight molecule to the molecular weightof the lowest molecular weight molecule. Polydispersity may becontrolled by carefully maintaining constant monomer concentrations andisothermal conditions within the reaction mixture. Generally speaking,it is desirable that the polydispersity be as low as possible in orderto diminish the content of unwanted relatively low or high molecularweight polyisobutylenes in the product and thus improve the quality ofthe latter. By following the concepts and principles of the presentinvention, it has been found that the polydispersity of the productmaybe controlled at no more than about 2.0. Preferably, through the useof the invention, a polydispersity of no more than about 1.65 may beachieved. Even more desirably, the polydispersity may be controlled soas to be within the range of from about 1.3 to about 1.5.

[0073] The polyisobutylene product obtained through the use of thepresent invention should generally have a terminal (vinylidene)unsaturation content of at least about 70%. That is to say, at leastabout 70% of the double bonds remaining in the polymerized productshould preferably be in a terminal position. Ideally the vinylidenecontent should be no less than about 80% or even higher. However,vinylidene content is indirectly related to conversion rates. That is tosay, the higher the conversion rate, the lower the vinylidene content.Moreover, vinylidene content is directly related in the same way tomolecular weight. Accordingly, in each process a balance is requiredbetween molecular weight, conversion rate and vinylidene content.

EXAMPLE 1

[0074] Using the principles and concepts of the invention, a reactorsuch as the reactor illustrated in FIG. 1, was used to produce a lowmolecular weight, highly reactive polyisobutylene. The feedstock wasessentially the same as shown above in Table 1, and the coolantcirculated on the shell side of the reactor was a mixture of 35 weight %methanol and 65 weight % water. The inlet coolant temperature was 32° F.A 1:1 BF₃/methanol complex catalyst was used. All pertinent reactor dataand dimensions are set forth below in Table 4. TABLE 4 Feedstock flowrate 1.7 gpm Recirculation flow rate 50 gpm Feedstock density 5 lb/galConversion 63 wt % Concentration of isobutylene in 92 wt % feedstockΔH_(reaction) 398 Btu/lb μ reaction mixture 4.5 cP = 0.0030 lb/ft sec Cpof reaction mixture 0.46 Btu/lb ° F. Reaction effective density 44.9lb/ft³ Thermal conductivity 0.075 Btu/hr ft ° F. Total volume of reactorrecirculation 390.2 in³ system Residence time 59.6 seconds Linearvelocity inside tubes 9.3 ft/sec Reynolds number 3180 Surface area oftubes 23.6 ft² Heat generated 1961 Btu/min ΔT_(lm) 37.3° F. Heat flux83.2 Btu/min ft² U 133.7 Btu/min ft² ° F. Cp of coolant 0.86 Btu/lb ° F.Density of coolant 7.70 lb/gal Coolant flow rate 39.3 gpm ΔT coolant8.0° F. Heat removed 2074 Btu/min

[0075] The composition of the product thus obtained is as set forthbelow in Table 5. TABLE 5 Crude Polyisobutylene Product IngredientWeight % C₃ components 0.00 I-butane 6.41 n-butane 1.68 1-butene 1.30I-butene 33.00 trans-2-butene 0.83 cis-2-butene 0.38 1,3-butadiene 0.21polyisobutylene 56.19

[0076] Again it is to be noted that one of the main objectives inaccordance with the invention is to provide a flow rate through thereactor and other parameters such that the reaction mixture is in agenerally constant state of turbulent flow during the reaction.Turbulent flow results in a twofold augmentation of the overall process.First, turbulent flow results in intimate intermixing of the contents ofthe reactor to enhance the kinetics of the reaction. Second, turbulentflow results in an enhancement of the tube side heat transfercoefficient to thereby improve the removal of the heat of the reaction.These results may be achieved by conducting the reaction on the tubeside of a shell-and-tube heat exchanger reactor and circulating acoolant on the shell side.

[0077] The foregoing description concerns methodology which permits thePIB polymerization reaction to be conducted at higher temperatures andat lower residence times than current processes. In accordance with thisembodiment of the present invention, a stable BF₃ catalyst system(BF₃/methanol) may be used. Moreover, an improved turbulent loop reactorconfiguration including a heat exchanger to effect simultaneous heatremoval is advantageously employed. The turbulent flow also enablesintimate mixing of the two-phase reaction system.

[0078] In addition to highly reactive PIB, the process of the inventionprovides an improved process for preparing oligomers and highermolecular weight polymers from olefinic precursors. In general, theprocess of the invention may be used to produce conventional PIB, lowmolecular weight oligomers of branched olefins, preferably isobutylene,oligomers and higher molecular weight polymers of linear C₃-C₁₅ alphaolefins, and oligomers and higher molecular weight polymers of C₄-C₁₅reactive non-alpha olefins. In accordance with this aspect of theinvention, and particularly where the desired product is a relativelylow molecular weight (<350 and perhaps even <250) oligomer, the catalystcomplex is desirably stable, even under the relatively higher reactiontemperatures needed for the production of oligomeric olefinic products.

[0079] Examples of processes for production of relatively low molecularweight oligomers of olefinic monomeric components are set forth below.In these examples, a loop reactor as illustrated in FIG. 2 is utilizedadvantageously. As illustrated in FIG. 2, the reactor 100 may consist ofa single tube 102 surrounded by a heat exchanger shell 104. In all otheressential aspects, the recirculation system may preferably be the sameas described in connection with the reactor 10 of FIG. 1, except that arecirculation line 106 is provided to return the recirculating residualmixture from the top of reactor tube 102 to the pump suction line 20.The exit line 55 is connected directly to recirculation line 106 asshown.

EXAMPLE 2

[0080] A stream containing 2.19 wt % Isobutane, 61.5 wt % n-butane, 0.64wt % 1-butene, 28.18 wt % trans-2-butene and 7.49 wt % cis-2-butene(35.66 wt % 2-butene) is introduced into the a loop reactor system ofFIG. 2 via feed line 15 at a rate of 156 ml/min (93.6 g/min). A catalystcomplex containing BF3/methanol complex (one mole of BF3 to one mole ofmethanol) is fed to the reactor at a rate of 8 ml/min (10.4 g/min). Thereaction temperature is maintained constant at 90° F. All pertinentreactor data and dimensions are set forth below in Table 6. The reactoreffluent exits the top of the reaction loop via line 55 and is fed intoa decant (not shown) were the catalyst is preferably separated out fromthe organic layer. A portion of the catalyst may then be recycled backto the reactor lowering the amount of fresh catalyst required. Theproduct coming out of the decant overhead is mixed with NH₄OH to quenchany remaining catalyst in the organic phase and is sent to a seconddecant. The products is washed twice more with water and decanted toremove the last traces of catalyst. The oligomer product composition isgiven in Table 7. TABLE 6 HC flow rate 0.0412 gpm Pump around flow rate1.5 gpm HC density 5 lb/gal % Conversion 51 wt % % 2-butene in feedstock 36.55 wt % ΔH_(rxn) 318 Btu/lb μ 0.6 cP = 0.0004 lb/ft-s Cp 0.46Btu/lb- ° F. Reactor OD 0.375 in Reactor wall thickness 0.035 in ReactorID 0.305 in Reactor length 10.5 ft Reactor volume 9.2 in³ # of tubes 1 #of passes 1 Residence time 58.02 seconds Linear velocity 6.59 ft/sSurface area 1.03 ft² Heat generated 12.2 Btu/min ΔT_(lm) 3.0° F. Heatflux 11.8 Btu/min-ft² U 237.0 Btu/hr-ft²- ° F. Re 15531

[0081] TABLE 7 C₈  7.9 wt % C₁₂ 29.8 wt % C₁₆ 35.9 wt % C₂₀ 16.1 wt %C₂₀₊ 10.3 wt %

EXAMPLE 3

[0082] A stream containing 94.0 wt % 1-decene and 6.0 wt % C₁₀-isomerswas fed into the loop reactor of FIG. 2 at a rate of 10 ml/min (7.4g/min). A catalyst complex containing BF3/methanol complex (one mole ofBF3 to one mole of methanol) was fed to the reactor at a rate of 1ml/min (1.3 g/min). The reaction was held at a constant temperature of70° F. All pertinent reactor data and dimensions are given in Table 8.Both the reactor setup and downstream catalyst removal steps areidentical to Example 2. The product stream contained about 59.8 wt % ofC₂₀ oligomers and about 40.2 wt % of C₃₀ oligomers. TABLE 8 HC flow rate0.00264 gpm Pump around flow rate 1.5 gpm HC density 6.2 lb/gal %Conversion 90 wt % % 1-decene in feed stock 94 wt % ΔH_(rxn) 318 Btu/lbμ 1.2 cP = 0.0008 lb/ft s Cp 0.50 Btu/lb ° F. Reactor OD 0.375 inReactor wall thickness 0.035 in Reactor ID 0.305 in Reactor length 10.5ft Reactor volume 9.2 in³ # of tubes 1 # of passes 1 Residence time905.13 seconds Linear velocity 6.59 ft/s Surface area 1.03 ft² Heatgenerated 4.4 Btu/min ΔT_(lm) 1.2° F. Heat flux 4.3 Btu/min-ft² U 213.2Btu/hr-ft²- ° F. Re 9604.4

[0083] As can be seen from the foregoing examples, the inventionprovides a process for preparing a polyolefin product having preselectedproperties. In accordance with the invention, the process advantageouslyemploys a stable complex of BF₃ and a complexing agent therefor. Theresidual reaction mixture in the reaction zone is recirculated at arecirculation rate sufficient to cause intimate intermixing of thereaction mixture. The rate is also such that the heat of reaction isremoved from the reaction mixture at a rate calculated to provide asubstantially constant reaction temperature in the reaction mixturewhile the same is recirculating in the reaction zone. The introductionof the feedstock and the withdrawal of the product stream arecontrolling such that the residence time of the olefinic componentsundergoing polymerization in the reaction zone is appropriate forproduction of the desired polyolefin product.

[0084] Although the foregoing text and examples have focused onprocesses wherein a single monomer is included in the feedstock, it willbe apparent to the routineers in the olefin polymerization art that inaccordance with the principles and concepts of the present invention,the feedstock may desirably, at times, include two or more monomers soas to produce useful copolymeric products.

We claim:
 1. A liquid phase polymerization process for preparing apolyolefin product having preselected properties, said processcomprising: providing a liquid feedstock comprising at least oneolefinic component; providing a catalyst composition comprising a stablecomplex of BF₃ and a complexing agent therefor; introducing saidfeedstock and said catalyst composition into a residual reaction mixturein a loop reactor reaction zone; recirculating the residual reactionmixture in said zone at a recirculation rate sufficient to causeintimate intermixing of the residual reaction mixture, the feedstock andthe catalyst composition to thereby present a recirculating, intimatelyintermixed reaction admixture in said reaction zone; maintaining therecirculating intimately intermixed reaction admixture in its intimatelyintermixed condition and removing heat of reaction from the reactionadmixture at a rate calculated to provide a substantially constantreaction temperature in the reaction admixture while the same isrecirculating in said reaction zone, said constant reaction temperaturebeing at a level appropriate for causing olefinic components introducedin said feedstock to undergo polymerization to form said polyolefinproduct in the presence of said catalyst composition; withdrawing aproduct stream comprising said polyolefin product from said reactionzone; and controlling the introduction of said feedstock into saidreaction zone and the withdrawal of said product stream from thereaction zone such that the residence time of the olefinic componentsundergoing polymerization in the reaction zone is appropriate forproduction of said polyolefin product.
 2. A process as set forth inclaim 1, wherein the reaction zone comprises a tube side of ashell-and-tube heat exchanger and said heat of reaction is removedsimultaneously with its generation by circulation of a coolant in theshell side of the exchanger.
 3. A process as set forth in claim 1,wherein said residence time is no greater than about 3 minutes.
 4. Aprocess as set forth in claim 1, wherein said residence time is nogreater than about 2 minutes.
 5. A process as set forth in claim 1,wherein said residence time is no greater than about 1 minute.
 6. Aprocess as set forth in claim 1, wherein said residence time is lessthan 1 minute.
 7. A process as set forth in claim 1, where saidcomplexing agent comprises an alcohol.
 8. A process as set forth inclaim 7, where said complexing agent comprises a primary alcohol.
 9. Aprocess as set forth in claim 8, where said complexing agent comprises aC₁-C₈ primary alcohol.
 10. A process as set forth in claim 7, where saidalcohol has no hydrogen atom on a β carbon.
 11. A process as set forthin claim 10, where said alcohol comprises methanol.
 12. A process as setforth in claim 10, where said alcohol comprises neopentanol.
 13. Aprocess as set forth in claim 1, where said complexing agent provides acomplex with BF₃ which is stable at temperatures needed to produceoligomeric olefinic products.
 14. A process as set forth in claim 1,where said complexing agent comprises a glycol.
 15. A process as setforth in claim 14, where said complexing agent comprises ethyleneglycol.
 16. A process as set forth in claim 1, wherein the molar ratioof BF₃ to complexing agent in said complex ranges from approximately0.5:1 to approximately 5:1.
 17. A process as set forth in claim 1,wherein the molar ratio of BF₃ to complexing agent in said complexranges from approximately 0.5:1 to approximately 2:1.
 18. A process asset forth in claim 1, wherein the molar ratio of BF₃ to complexing agentin said complex ranges from approximately 0.5:1 to approximately 1:1.19. A process as set forth in claim 1, wherein the molar ratio of BF₃ tocomplexing agent in said complex is approximately 1:1.
 20. A process asset forth in claim 1, wherein from about 0.1 to about 10 millimoles ofBF₃ are introduced into said reaction admixture with said catalystcomposition for each mole of olefinic component introduced into saidadmixture in said feedstock.
 21. A process as set forth in claim 1,wherein from about 0.5 to about 2 millimoles of BF₃ are introduced intosaid reaction admixture with said catalyst composition for each mole ofolefinic component introduced into said admixture in said feedstock. 22.A process as set forth in claim 1, wherein the reaction admixture isrecirculated at a first volumetric flow rate, and said feedstock andsaid catalyst composition are introduced at a combined second volumetricflow rate.
 23. A process as set forth in claim 22, wherein the ratio ofsaid first volumetric flow rate to said second volumetric flow rate issuch that the concentrations of ingredients in the reaction admixtureremain essentially constant.
 24. A process as set forth in claim 22,wherein the ratio of said first volumetric flow rate to said secondvolumetric flow rate is such that essentially isothermal conditions areestablished and maintained in said reaction admixture.
 25. A process asset forth in claim 22, wherein said feedstock and said catalystcomposition are premixed and introduced into the reaction zone togetheras a single stream at said second volumetric flow rate.
 26. A process asset forth in claim 22, wherein said feedstock and said catalystcomposition are introduced into the reaction zone separately as twostreams, the flow rates of which together add up to said secondvolumetric flow rate.
 27. A process as set forth in claim 22, whereinthe reactor configuration, the properties of the reaction mixture, andthe first volumetric flow rate are such that turbulent flow ismaintained in said reaction zone.
 28. A process as set forth in claim27, wherein the reaction zone is a tube side of a shell-and-tube heatexchanger.
 29. A process as set forth in claim 28, wherein a U of atleast about 50 Btu/min ft² ° F. is maintained in said reaction zone. 30.A process as set forth in claim 1, wherein said feed stock comprises atleast about 30% by weight of said olefinic component.
 31. A process asset forth in claim 1, wherein said feed stock comprises non-reactivehydrocarbon diluents.
 32. A process as set forth in claim 31, whereinsaid feed stock comprises at least about 30% by weight of said olefiniccomponent with the remainder being said diluents.
 33. A process as setforth in claim 22, wherein said reaction zone is the tube side of ashell-and-tube heat exchanger.
 34. A process as set forth in claim 1,wherein the molar ratio of BF₃ to complexing agent in said complex isapproximately 0.75:1.
 35. A process as set forth in claim 1, wherein thepolymerization process is a cationic process.
 36. A process as set forthin claim 1, wherein the polymerization process is a covalent process.37. A process as set forth in claim 1, wherein said polyolefin producthas a molecular weight no more than about
 5000. 38. A process as setforth in claim 37, wherein said polyolefin product has a molecularweight of at least about
 350. 39. A process as set forth in claim 38,wherein the olefinic component comprises isobutylene and the polyolefinproduct comprises polyisobutylene.
 40. A process as set forth in claim39, wherein said polyisobutylene has a vinylidene content of at leastabout 50%.
 41. A process as set forth in claim 10, wherein the olefiniccomponent comprises a C₃ to C₁₅ linear alpha olefin.
 42. A process asset forth in claim 14, wherein the olefinic component comprises a C₃ toC₁₅ linear alpha olefin.
 43. A process as set forth in claim 10, whereinthe olefinic component comprises a C₄ to C₁₅ reactive non-alpha olefin.44. A process as set forth in claim 43 wherein the olefinic component is2-butene.
 45. A process as set forth in claim 14, wherein the olefiniccomponent comprises a C₄ to C₁₅ reactive non-alpha olefin.
 46. A processas set forth in claim 45 wherein the olefinic component is 2-butene. 47.A process as set forth in claim 37, wherein said polyolefin product hasa molecular weight of at least about
 250. 48. A process as set forth inclaim 1, wherein said liquid feedstock comprises a raff-1 stream.
 49. Aprocess as set forth in claim 48, wherein the olefinic componentcomprises isobutylene and the polyolefin product comprisespolyisobutylene.
 50. A process as set forth in claim 39, wherein saidliquid feedstock comprises a raff-1 stream.
 51. A liquid phasepolymerization process for preparing a polyolefin product havingpreselected properties, said process comprising: providing a liquidfeedstock comprising at least one olefinic component; providing acatalyst composition comprising a stable complex of BF₃ and a complexingagent therefor; introducing said feedstock and said catalyst compositioninto a residual reaction mixture in a loop reactor reaction zone;recirculating the residual reaction mixture in said zone at arecirculation rate sufficient to cause intimate intermixing of theresidual reaction mixture, the feedstock and the catalyst composition tothereby present a recirculating, intimately intermixed reactionadmixture in said reaction zone; maintaining the recirculatingintimately intermixed reaction admixture in its intimately intermixedcondition and removing heat of reaction from the reaction admixture at arate calculated to provide a substantially constant reaction temperaturein the reaction admixture while the same is recirculating in saidreaction zone, said constant reaction temperature being at a levelappropriate for causing olefinic components introduced in said feedstockto undergo polymerization to form said polyolefin product in thepresence of said catalyst composition; withdrawing a product streamcomprising said polyolefin product from said reaction zone; andcontrolling the introduction of said feedstock into said reaction zoneand the withdrawal of said product stream from the reaction zone suchthat the residence time of the olefinic components undergoingpolymerization in the reaction zone is appropriate for production ofsaid polyolefin product, wherein the reaction admixture is recirculatedat a first volumetric flow rate, and said feedstock and said catalystcomposition are introduced at a combined second volumetric flow rate,and wherein the ratio of said first volumetric flow rate to said secondvolumetric flow rate ranges from about 20:1 to about 50:1.
 52. A processas set forth in claim 51, wherein the ratio of said first volumetricflow rate to said second volumetric flow rate ranges from about 25:1 toabout 40:1.
 53. A process as set forth in claim 52, wherein the ratio ofsaid first volumetric flow rate to said second volumetric flow rateranges from about 28:1 to about 35:1.
 54. A liquid phase polymerizationprocess for preparing a polyolefin product having preselectedproperties, said process comprising: providing a liquid feedstockcomprising at least one olefinic component; providing a catalystcomposition comprising a stable complex of BF₃ and a complexing agenttherefor; introducing said feedstock and said catalyst composition intoa residual reaction mixture in a loop reactor reaction zone;recirculating the residual reaction mixture in said zone at arecirculation rate sufficient to cause intimate intermixing of theresidual reaction mixture, the feedstock and the catalyst composition tothereby present a recirculating, intimately intermixed reactionadmixture in said reaction zone; maintaining the recirculatingintimately intermixed reaction admixture in its intimately intermixedcondition and removing heat of reaction from the reaction admixture at arate calculated to provide a substantially constant reaction temperaturein the reaction admixture while the same is recirculating in saidreaction zone, said constant reaction temperature being at a levelappropriate for causing olefinic components introduced in said feedstockto undergo polymerization to form said polyolefin product in thepresence of said catalyst composition; withdrawing a product streamcomprising said polyolefin product from said reaction zone; andcontrolling the introduction of said feedstock into said reaction zoneand the withdrawal of said product stream from the reaction zone suchthat the residence time of the olefinic components undergoingpolymerization in the reaction zone is appropriate for production ofsaid polyolefin product, wherein the reaction admixture is recirculatedat a first volumetric flow rate, wherein the reactor configuration, theproperties of the reaction mixture, and the first volumetric flow rateare such that turbulent flow is maintained in said reaction zone, andwherein a Reynolds number of at least about 2000 is maintained in saidreaction zone.
 55. A process as set forth in claim 54, wherein thereactor is a tube side of a shell-and-tube heat exchanger, and a heattransfer coefficient of at least about 50 Btu/min f² ° F. is maintainedin said reaction zone.
 56. A liquid phase polymerization process forpreparing a polyolefin product having preselected properties, saidprocess comprising: providing a liquid feedstock comprising at least oneolefinic component; providing a catalyst composition comprising a stablecomplex of BF₃ and methanol; introducing said feedstock and saidcatalyst composition into a residual reaction mixture in a loop reactorreaction zone; recirculating the residual reaction mixture in said zoneat a recirculation rate sufficient to cause intimate intermixing of theresidual reaction mixture, the feedstock and the catalyst composition tothereby present a recirculating, intimately intermixed reactionadmixture in said reaction zone; maintaining the recirculatingintimately intermixed reaction admixture in its intimately intermixedcondition and removing heat of reaction from the reaction admixture at arate calculated to provide a substantially constant reaction temperaturein the reaction admixture while the same is recirculating in saidreaction zone, said constant reaction temperature being at a levelappropriate for causing olefinic components introduced in said feedstockto undergo polymerization to form said polyolefin product in thepresence of said catalyst composition; withdrawing a product streamcomprising said polyolefin product from said reaction zone; andcontrolling the introduction of said feedstock into said reaction zoneand the withdrawal of said product stream from the reaction zone suchthat the residence time of the olefinic components undergoingpolymerization in the reaction zone is appropriate for production ofsaid polyolefin product.
 57. A single stage liquid phase polymerizationprocess for preparing a polyolefin product having preselectedproperties, said process comprising: providing a liquid feedstockcomprising at least one olefinic component; providing a catalystcomposition comprising a stable complex of BF₃ and a complexing agenttherefor; introducing said feedstock and said catalyst composition intoa residual reaction mixture in a loop reactor reaction zone, saidreaction zone being the only stage of the process where substantialpolymerization occurs; recirculating the residual reaction mixture insaid zone at a recirculation rate sufficient to cause intimateintermixing of the residual reaction mixture, the feedstock and thecatalyst composition to thereby present a recirculating, intimatelyintermixed reaction admixture in said reaction zone; maintaining therecirculating intimately intermixed reaction admixture in its intimatelyintermixed condition and removing heat of reaction from the reactionadmixture at a rate calculated to provide a substantially constantreaction temperature in the reaction admixture while the same isrecirculating in said reaction zone, said constant reaction temperaturebeing at a level appropriate for causing olefinic components introducedin said feedstock to undergo polymerization to form said polyolefinproduct in the presence of said catalyst composition; withdrawing aproduct stream comprising said polyolefin product from said reactionzone; and controlling the introduction of said feedstock into saidreaction zone and the withdrawal of said product stream from thereaction zone such that the residence time of the olefinic componentsundergoing polymerization in the reaction zone is appropriate forproduction of said polyolefin product.
 58. A process as set forth inclaim 57, wherein said complexing agent comprises methanol.